Catalytic hydrogenation of dripolene



United States Patent Ufiiice 2,953,612 Patented Sept. 20, 196i.)

CATALYTIC HYDROGENATION F DRIPOLENE Mauford R. Haxton, Texas City,Walker F. Johnston, IL, La Marque, and Irvin F. Teykl, Houston, Tex.,assignors to The American Oil Company, Texas City, Tex., a corporationof Texas Filed Mar. 3, 1958,.Ser. No. 718,577

9 Claims. (Cl. 260683.9)

This invention relates to the catalytic hydrogenation of olefinichydrocarbons. More specifically the invention is concerned with animproved process for-the catalytic hydrogenation of dripolenefractions,dripolene being the normally liquid mixture of hydrocarbons obtained asa byproduct in the high temperature pyrolysis of gaseous hydrocarbons.

It is well known that the hightemperature pyrolysis of gaseoushydrocarbons to prepare ethylene results in the by-production of anormally liquid mixture of hydrocarbons through reactions such-aspolymerization, alkylation, aromatization, dehydrogenation, and thelike. The mixture is commonly termed. dripolene, and While it containsvirtually all classes of hydrocarbons it'predorninates in olefinic andaromatic compounds, mainly benzene. As the demand for ethylene for theproduction of polyethylene and plastics and'other petrochemicals rises,increasingly large supplies of dripolene are becoming available. Becauseof its benzene content, dripolene represents an exceedingly valuablematerial, and one which is steadily becoming more obtainable.

Thus far three large-volume usages have developed for dripolene. It isblended into motor gasolines, where the high octane numbers of itsaromatic and olefinic components render dripolene a desirable blendingstock. Dripolene may be fed to aromatics extraction units for therecovery of pure benzene. Finally, exceedingly valu-' able resins havebeen made by thermal or catalytic polymerization of high boilingdripolene fractions.

Dripolene however contains cyclic diolefins, and these give rise toproblems which have somewhat restricted the quantity of dripoleneblended into motor fuels or fed to aromatics extraction units. Cyclicdiolefins tend to form gum-like polymers in the presence of air, orupon-heating and for this reason only relatively small amounts ofdripolene can be blended into premium motor fuels. And in aromaticsextraction units, it is found that cyclic diolefins tend to concentratein the aromatics extract, thereby complicating the preparation of purearomatic compounds.

It has previously been disclosed in Oil audGas Journal, volume 52, May11, 1953, page 124, and in Haensel US. Patent Number 2,799,627 that a130-360 F. ASTM distillation boiling range dripolene fraction maybetreated in'the presence of a platinum-alumina-halogen catalyst atelevated temperatures and pressures andin the pres ence of hydrogen gasto selectively hydrogenate mo'noole fins and diolefins withoutsimultaneously hydrogenating the aromatic compounds to naphthenes.However, in attempting to utilize this process, We have found that aninordinate amount of a coke-like material forms throughout'the chargestock preheater system and throughout the catalyst bed. In fact, withthis process, after on-stream times of less than a week a solidcoke-like matrix: forms in the. reactors which completely preventsfurther operations and necessitates shut-downs for coke and catalystremoval. It is therefore an objectoftthe present-invention to provide animproved method'of conducting the 2 catalytic hydrogenation of olefinichydrocarbons in dripo= lene fractions which enables the-process tobe'conducted for periods substantially in excess of those heretoforeobtainable.

-In accordance with the object above, we have now discovered a method ofsubstantially reducing the formation of troublesome coke-like deposits.Briefly, we have discovered that coke formation is,virtually eliminatedif the dripolene fraction and a hydrogen-containing gas stream initiallycontact the platinum-alumina hydrogenation catalyst at conditions oftemperature and pressure such-that substantially all (i.e. about ormore, and preferably; at least of the dripolene remains in the liquidphase, while maintaining a critically low hydrogen sulfide concentrationin the hydrogen-containing gas stream. By means of our discovery it ispossible to conduct the bye drogenation of dripolene in a convenientmanner and with on-stream times of many weeks or monthswithout the needfor catalyst regeneration. Our system maybe embodied in either oftwo-variations, each of which has advantages under particularcircumstances. In the first embodiment only one reaction zone isemployed, and the commingled hydrogen-containing gas and liquid chargestock stream contacts a substantially adiabatic catalyst zone, and theexothermic heat of hydrogenationserves to increase the commingledstreamtemperature as it passes through the zone. In this embodiment the amountof catalyst, expressed as weight hourly space velocity (weight of chargestock per unit weight of catalyst per hour) is at least effective tohydrogenate olefinic hydrocarbons without the need for application ofexternal heat. In the: second embodiment, two'reaction' zonesareemployed, with interstage heating between the two. In the first zone,the dripolene charge stock is substantially in the liquid phase whenthecommingled streamof charge stock and hydrogen-containing gas contact thecatalyst, and after partial hydrogenation an interstage heater raisesthe temperature between reaction zones in order to provide a highertemperature stream to the second zone and thus minimize the total amountof catalyst necessary. It will be noted that the single-zone embodimentrequires no external heating facilities.

The amount of hydrogen sulfide thatmay be present in our process isquite critical, andaccordingly we find it essential to provide tothereaction zone a hydrogen-con taining gas stream containing less than 12grains of hydrogen sulfide per standard cubic feet of gas. If: forexample the level of sulfur in the hydrogen-containing gas exceeds 12grains, both the catalyst activity and the catalyst life diminishrapidly, and if the level increases to as much as 57 grains per 100s.c.f., the catalyst-becomes completely deactivated in a matter ofminutes. Fortunately however, the eiiect of either hydrogen sulfide gasor mercaptan sulfur in the' dripolene feed on-the hydrogenation processappears to be temporary'with respect to its effect on product qualityalthough periods of high sulfur do materially increase the amount ofcoke deposition.

Suitable platinum-alumina hydrogenation catalysts are conveniently thosecatalysts which have been found eminently suitable for use in naphthahydroforming-processes. Generally, these catalysts containfromabout=0.0l to about 10% by weight of platinum and may optionallyfrom about 0.05 to about 3% by Weight of a halogen, preferably chlorineand/or fluorine, on a high surface area alumina support such as thealumina described in Heard Reissue Patent Number 22,196. The catalystsmay bein the form of pills, pellets, extrudates, spheres or the like,and conventionally have a size between about since it appears thatcatalysts previously used for hydroforming are more stable and have lesstendency to hydrogenate aromatic compounds than fresh platinumaluminacatalysts. Furthermore, used catalysts exhibit less tendency to causewasteful hydrocracking of hydrocarbons and thus result in higher yieldsof recoverable liquid product.

The conditions of pressure, temperature, liquid hourly space velocityand hydrogen-containing gas rate which are employed in the practice ofthe present invention are inter-related such that the commingledfeedstock and hydrogen-containing gas, as it initially contacts thecatalyst bed, consists of a gas phase and a liquid phase wherein theliquid phase comprises substantially all of the charge stock. It isdesired that at least 80 mol percent, and preferably at least 90 molpercent of the dripolene charge contact the catalyst as a liquid.Pressures within the range of 100 to 1000 p.s.i.g. are desired, withpressures from 300 to 500 p.s.i.g. preferred from a comercial standpointas this latter range favors conditions at which the hydrogenationreaction occurs rapidly. Within the broad pressure range the bed inlettemperature may be between 50 and 200 F., most desirably between 100 and150 F., typically 115 F. With most dripolene stocks the temperature risethrough an adiabatic bed, for complete olefinic saturation, is on theorder of 350450 F. and provides an average reactor temperature of about280340 F. This average temperature may be increased by providing morecatalyst or may be decreased by increasing the proportion ofhydrogen-containing gas to charge stock. With respect to thehydrogen-containing gas, it is desirably employed in a proportion of 500to 10,000 standard cubic feet per barrel of charge stock, preferablyfrom 1000 to 4000 s.c.f./b., e.g. 1500 s.c.f./b. This gas preferably iscomposed of at least 70% hydrogen aud may be derived from a naphthahydroforming operation. Although the experimentally observed consumptionof hydrogen varies between 600 and about 650 s.c.f./b., it is preferredto maintain a substantially larger amount in the reaction zone. This maybe accomplished economically by recycling the excess hydrogen. Thehydrogen-containing gas, if recycled, must be chemically treated tomaintain the critically low hydrogen sulfide level therein.

Dripolene employed as the charging stock in our invention is a portionof the hydrocarbon liquid obtained by the high temperature pyrolysis ofa normally gaseous hydrocarbon containing at least two carbon atoms inthe molecule, or a mixture of such hydrocarbons. The normally gaseoushydrocarbon which is charged to the high temperature pyrolysis may be abyproduct refinery gas. In the preparation of dripolene, a gaseoushydrocarbon such as ethane, propane, propylene or a mixture of suchhydrocarbons is preheated and passed through an alloy tube at a highspace velocity and a pyrolysis temperature between about 1200 and 1800R, preferably between about 1350 and 1550 F. Low pressures up to about100 p.s.i.a. are ordinarily employed in this operation, a pressure belowabout 35 p.s.i.-a. being satisfactory. The time of exposure to the hightemperatures is usually about 0.05 to 5 seconds, contact times of 0.1 to1 second being prefered.

The pyrolysis produces normally gaseous products containing unsaturatedhydrocarbon such as ethylene, normally liquid hydrocarbons rich inunsaturated hydrocarbons including olefins and diolefins of varyingboiling points and structural configuration, and various aromatichydrocarbons, as well as tar. The unsaturated hydrocarbons such asethylene which are contained in the normally gaseous product are usuallythe desired product of the pyrolysis process. The normally liquidhydrocarbons and tar which are obtained are considered to be byproductsof the pyrolysis operation. High temperature pyrolysis products arerapidly cooled, usually by quenching with water to a temperature ofabout 400 F. A viscous tarry material condenses out of the gas duringthe quenching. The gases from the quenching operation are compressed andcooled and a liquid material which boils between about 100 and 400 F.condenses out of the gases during the compressing-cooling step. Thisliquid is dripolene. The amount of tar and dripolene produced isdependent upon the feed, temperature, contact time and pressure. Thequantity of liquid hydrocarbons produced in this way is ordinarily about3% by weight of the total quantity of gas charged to the pyrolysisreactor.

The normally liquid mixtures of hydrocarbons which is termed dripolenehas never been completely analyzed because of its complexity. A typicalspecimen of total dripolene was characterized as follows:

Index of refraction, 11 1.4830

A typical example of dripolene was analyzed by conventional techniquesand was found to contain the following compounds in the amountsspecified:

TABLE II Analysis, volume-percent:

Propane and propylene 0.7 Isobutane 0.1 "Isobutylene 0.8 l-butene 0.52-b-utene 0.6 n-butane 0.4 Butadiene 3.9 Pentadienes 7.7 Pentylenes 6.3Other C a 0.4 Benzene 34.2 Toluene 7.8 Xylenes 1 Styrene 3Dicyclopentadiene 5 Other 29.6

Our invention is particularly concerned with the fraction of dripoleneboiling Within the range of about 100- 375 F., although it is notessential that the dripolene boil entirely within the range or that allof the dripolene fraction boil within the range. Our charging stock isobtained as an overhead or heartcut in the distillation of totaldripolene to obtain about 70 to 90% of charge fraction, While thebottoms may be used to prepare resins by processes well known to theart.

To more fully describe the process of our invention and to illustratethe two embodiments thereof, attention is directed to the attachedflowsheet showing a hydrogenation unit designed and adapted to producefrom dripolene either an aromatic concentrate or a high octane motorfuel blending stock by either embodiment.

Dripolene liquid is withdrawn from external storage tanks and conductedthrough line 1 to fractionator 2 which is provided with corrosionresistant distillation trays or perforated pans, wherein an overheaddripolene charge stock fraction comprising about of the total dripoleneis separated by distillation. from about.-20%-of high boiling bottoms,which latter is sent via-line 3 to the: resins plant, not shown.v Thetotal dripolene fed to fractionator 2 has an analysisapproximating thetypical dripolene described previously. .The 80% 'fractionator 2overhead which is taken through line 4 has the following composition:

TABLEIII Charge analysis:

Gravity, API 32.4 RVP, p'.s.i.a 6.8 ASTMdistillation, 'F.:

IBP 134 10% 163 30% 1 79 50% 189 70% 202 90% 280 FBP 356:Lightlhydrocarbvnsanalysis- Component:

C liquid volzrpercentm. -.1 i634: do' 053 1C4: do .02 263;: do 0.2 3 -nCa ido= .012 C; dioiefirr do 1116 C diolefin. .i ;db 61.6 C monoolefin do3 .1 C paraffin do 0.2 h C do 87.5 B'enzene' vol.-percent 53 (36gravity, API 28.0

The 80% dripolene charge fraction contains 70 parts per million sulfurand 29parts'per million organic' chlor= ides, and has a bromine numberof48f (indicative of total olefins) and a maleic anhydride value (MAV,representing conjugated diolefins) of 47- m-g./g." The bottomswithdrawnthrough line 3 has an ASTM distillation boil- -range' betweenabout 200 -and--400--F., preferably between about 230 and375" F.

The'dripolene charge" is' conducted'through' line '4; cooler 5, and line6 to charge pump 7', which-may. be a multistage centrifugali pumpadaptedto pump the d'ripo leneohargeto the reactorsystem operating at apressure Qf3'2'5f p61111ds-p8r squarein'chtgag'e'. "The cooler 5outlettemperature is: about 80 The charge-stockfrom pump? is sentthrough'linej' 8 tofjuiijcture'g; where n-1 is metby a stream of recyclehydrogen containing gasffrom :line I0"in1' the, amount of I350" standardcubic feet of total hydrogen-containing gasperbarrel- .of charge. -The:gas has aic'ompo'si'ti'on of approximately 7 80% hydrogen, with thebalanee-consistingprimaril-y of-methan'e, ethane, and some, propaneiandpropylene, together, with less than the critical" limitof 12 grains of HS--per 100 standard cubic feetQofit otaI. gas. .Itlisl highly Ipreferrdlt'hatifthis :gas contain; if possible; loss than- 3 grains per- 100:cubic feet of I-ll sii'lhetemperatur e .ofltlielcomniin'gledi liquid andgas stream is 1 1 5 F.,--and-at this temperature the -'oonrmihgledstreampasses into-reaction as employed. and :is, shown. symbolically asa single bed orjchamberl'l, although: it may comprise 'a :pl:ura.-lityof :serially' ;or parallel-connected reactionchambers. At theseoperating conditions; 94 mol percent of the dripolene is invthe liquid:phase when: the commingled stream; initially con- .fatts-theccatalyst.

I=n1the first embodiment, a single reactionszonea reaction zone- .11',operates essentially adiabaticallly, @that is the .commingled dripo'lenecharge and hydrogenecontaining :gas :stream are permitted to'increase intemperatureby the: exothermic. heat of monoolefin and. diolefin hyd-ro:genationon passage through the.- catalyst bed. The catalyst employed isspent, UI-traf rming catalyst obtained after more than one yearsuse. ina regenerative naphtha hydroforming unit and has-an activity forhydroforming of substantially less. than that of fresh Ultraforming.catalyst, but is. very. nearly as active for hydrogenation. as is freshcatalyst. The catalyst in Reaction Zone. I, i.e. chamber 11, is in theform of pellets having an average lengthand diameter approximating andiswdisposed so as to'permit downflow passage of the-commingled stream. Aweight hourlyspace velocityof '2 is" em"- ployed. In passage throughReaction; ZoneI' the dripoq lene plus hydrogen stream temperature isincreased to 625'? R, which provides an average reaction temperature of370 F. In this zone, 625 standard cubic feet per barrelof hydrogen is.consumed by olefin hydrogenation, a quantity whichcompares closelywiththe theoretical hydrogen consumption based on the observedexperimental heat of reaction, 280' B.t.u./1b. The: quantity of catalystin Reaction Zone I is that which providesa weight. hourly space velocityof 2.0, i.e. 2.0 poundsiof dripolene charged per hour for each pound ofcatalyst. in Zone-I. 'I-hehydrogenated stream leaving'chamber I1 passesthrough -line12, line15, valve 16, line 1-7, and line24 to-cooler 25,.andthen through line .26: to receiver 27. Valves 14 and 22 are closed,thus blanking cit. heater 18- and reactor20 which are not used.

-As an alternate embodiment to the use of a single Reaction Zone I, amodification may he employed wherein the necessary quantity of catalystcan be:reduced substantially. In this second embodiment, thecatalyst-"is loaded inL' adiabaticReaction Zone I to: provide a weighthourlyspa-ce velocity of from say 4 to 20' (less than half the loadingas in the previous embodiment), and-the product stream leaving chamber11 through line 1-2,finstead of: passing-"through. by-pass linel's, a'valve 16 and line". 17, passes through heater lilland a second adiabaticreaction Z0118,- ;Reacti'on Zone II. This latter zone isrepresented:symbolically bychamber 20 which also may be a plurality ofserially or parallel-connected reaction chambers. Reaction Zone IIalsohas a catalyst loading to provide? a; weight hourlyspace velocity offrom 4- to 2.0;. butnot'necessarily the same loading as in ZonezI.twoi-zon'e embodiment, only a fraction of the olefin hydrogenation iscompleted in Reaction-Zone I, andras a consequence the. partiallyhydrogenated: stream is-reheatedi by heater-1:8 after passage througlrline 12, valve: 1-4, and line 13, to: a temperature within the range ofzabout'200 to 500 F. Thereafter, it is-passed. through line 19 toReaction Zone II and thence via line 21:, valve 22, line 23, and line 24to .cooler 25 and receiver 27. For this operation, valve 16 is closed.Thus, by providing a-higher temperaturein Reaction. Zone II and conseequently amore. rapid reaction rate, olefin hydrogenation proceeds morerapidly with a consequently reduced catalyst requirement while stillretaining the benefits of the invention in having liquid phasehydrogenation occurring in 'Reaction"Zorre"I;"" 'Thus,'the cokedeposition of: prior art processes is substantially reduced" by' theelimination of .preheater's and by commencing hydrogenation while mostof'the d'ripo'l'enecha'rge' is in"the"liquid"phase.

.In either alternative, thehydrogenated product stream comprisingpartially hydrogenated drip'olene in vapor form together with excesshydrogen-containing gas is cooled'in cooler 25 wherein the hydrogenateddripolene fraction condenses as a liquid which is sent, along with thenoncondensiblehydrogen-containing gas, to receiver 27. At receiver27;the 1 hydrogen-containing gas is separated and withdrawn through line 28and conductedtvia line 31 to amine scrubber 32, where a descendingvstream of diethanolamine or other agent effective to absorb H S isemployed to remove hydrogen sulfide gas formed by the destructivehydrogenation of sulfur compounds in? the dripolene charge. The amineisWithdrawn from line 34 and heated in a stripper, not shown, for thepur'-. pose of releasing absorbed hydrogen sulfide,

7 Since hydrogenation results in a net consumption of hydrogen on theorder of 600-650 standard cubic feet per barrel, it is necessary toreplenish this by the addition of hydrogen from an external source,conveniently a naphtha hydroforming unit. Depending upon the pressureand hydrogen sulfide concentration of the hydroformerhydrogen-containing gas, it may be added at either valved line 30,valved line 40 or valved line 42. Briefly, if the hydroformer gas isrelatively high in hydrogen sulfide, irrespective of its pressure, it isconducted into the system through valved line 31 where it can pass intothe amine scrubber 32 for hydrogen sulfide removal. However, where thehydroformer gas is relatively low in hydrogen sulfide, it may be addedto the system either through valved line 40 if at low pressure or valvedline 42 if at high. The composition of hydroformer gas varies with theoperation of the hydroformer and may range for example from 70-95%hydrogen, the balance being saturated light hydrocarbons such asmethane, ethane and propane. If this gas is of a purity below about 80%it may be desirable to vent a portion of the gas from receiver 27through valved vent line 54 so as to prevent a build-up ofnoncondensible methane, ethane and propane within the recycle gassystem. Where large quantities of hydroformer gas are available, thepresent recycle gas system may be eliminated in favor of a oncethroughhydrogen flow.

After treatment in amine scrubber 32, the essentiallyhydrogen-sulfide-free hydrogen-containing gas (i.e. containing less than12 grains of H 8 per 100 standard cubic feet) is conducted via line 35to water scrubber 36 where a descending stream of water from line 37scrubs entrained or vaporized amine from the gas. The rich water streamis withdrawn through line 38 and is concentrated for amine recovery in adistillation column, not shown. If desired, water vapor removalfacilities such as a glycol scrubbing tower or a silica gel or aluminadrier may follow water scrubber 36 in line 39.

The treated gas passes from water scrubber 36 through line 39 to thesuction of recycle gas compressor 41. Compressor 41 recycles the gasthrough lines 43, 44 and back to the juncture 9 with dripolene chargeline 8 and thence to Reaction Zone I.

Returning now to receiver 27, hydrogenated dripolene as a liquidcondensate passes through line 29 to stabilizer 45. The unstabilizedhydrogenated dripolene, obtained from hydrogenation in a single ReactionZone I, has the following analysis:

TABLE IV Gravity, API 34.4 Reid vapor pressure, p.s.i 6.0

ASTM distillation, F.:

IBP 126 10% 162 30% 178 50% 188 70% 202 90% 296 FBP 412 F-l octane,clear (research) 100.2 F-2 octane, clear (motor) 88.5 C gravity, API33.0 C gravity, API 30.1

Light hydrocarbon analysis Component:

iC Liquid volume percent 0.1 nC do 2.3 iC do 1.4 nC do 4.8 C do 91.4Bromine No., cg./g. 5 MAV, mg./g. 0.2

The product yields by catalytic hydrogenation are set forth below. Theincrease in volume and in weight percent recoveries is primarily due toan increase in volume and weight produced by the hydrogenation. V

TABLE V Product Yields. 1 Yields,Wt. Yields, Vol.

Percent Percent C1 0. 2 C2. 0. 1 G2 0. 1 i0: 0. 1 2 120 1. 9 .8 1'04 1.3 1. 9 n6 3. 7 5.0 00+ 93. 7 92.1

Total 101. 1 102. 0 Berwene V V 53 Unstabilized Product 101. 4

1 Yields are based on dripolene overhead charged.

TABLE VI Properties of hydrogenated dripolene overhead fractions IBP-140F. Fraction 140-190 F. Fraction 190 F. plus Fraction Weight Percent ofCharge Composition of Fraction:

Total C; Olefins V Benzene In Fraction, V0 Percent. Aromaties, Vol.Percent Olefins, Vol. Percent. Saturates, Vol. Percent Corrosion,ASTMCopper Strlp. Doctor- RSH Acidity Thiophene S ur J at Sulfur ResearchOctane, Clear .5

Research Octane, +1.0 cc. TEL- 101. 6

Research Octane, +3.0 cc. TEL. 104. 3

RVP, p s i 0. 6

Gravity, AP 28. 2

ASTM Distillation:

IBP 230 10%-. 242 30%- 254 50%. 270 70%. 306 390 PB 410 pure benzenefraction it is desirable to pass the -190" F. fraction from line 49through heater 51, line 52, and clay treater '53 at about 50 p.s.i.gpressure and at the boiling point of the fraction at this pressure forthe purpose of polymerizing diolefins. Alternatively, sulfuric acid ormaleic anhydride may be used for this treatment.

aasaeta After diolefim removal the 1:40-1:90" F. fraction is. passedthrough line fl and-conducted to an aromatics extraction unit employingknown selectii e solvents such as diethylene glycol-water, :dieth yleneglycobtriethylene. glycolwater; phenol; sulfur dioxide, or P'ourex or asilica .gel chromatographic adsorbent;

- 'R'et'urning to stabilizer- 45, the- IBP-1'40 fraction containing 90%;.pentanes'i pentenes: is highly useful as a gasoline blending componentto provide front end volatility properties. "Similarly, the 190 F.-E.P.fraction takenas. a-bottoms..th1'.ough. line.50.splitter 48 is composedof about 78% aromatics, largely -boiling within the toluene and-xylenerange,.:and:as. shown has .an .exceedinglly valuablehighcctane:rating,imaking it .a. desirable blending: component fonmotorv fuels. Itsresearch octanesclear'is 98.5, and..with 3; ccs.;:ofitetraethyl lead pergallonais ".Thus' it iszefiidenhthatour process-is eminently/suitablefor hydrogenating olefinic hydrocarbons in a low boiling dripolenefraction. Of particular interest is the fact that benzene comprises 53volume percent of the dripolene charged to Reaction Zone I, and also is53 volume percent of the total liquid product obtained at receiver 27.This indicates that substantially none of the benzene is hydrogenatedunder our reaction conditions, although about 90% or more of themonoolefins' and the diolefins are saturated. The above run wasconducted for a total of 280 hours and during this time the maleicanhydride value of the total product remained below 2 while the brominenumber remained below 14.

While there exists a very important limit with respect to the maximumtolerable hydrogen sulfide level in the hydrogen-containing gas, it hasbeen found that organic sulfur compounds in the dripolene charge have afar lesser effect on the process and do not lead to coke formation. Forexample, when the normal 70 p.p.m. sulfur content of dripolene wasincreased four hundred fold by the addition of butyl mercaptan, theproduct maleic anhydride value remained unchanged. However at the sametime the product bromine number increased from 4 to 16. When butylmercaptan addition was discontinued the bromine number returned to 4.These indicate that organic sulfur does not affect hydrogenation ofdiolefins but does alter the catalysts activity for monoolefinhydrogenation. Thus when it is desired to employ the total hydrogenateddripolene as a motor fuel blendstock, organic sulfur compounds can beadded to the charge for the purpose of retaining high-octane monoolefinin the product while selectively hydrogenating polymer-formingdiolefins.

Where, however, the dripolene charge contains excessive sulfur, or whenit is desired to produce a very high purity product, mercaptan removalfacilities may be installed. Sodium or potassium hydroxide,caustic-methanol or similar extraction facilities, placed preferablybefore fractionator 2 can remove any dissolved hydrogen sulfide and mostof the lower molecular weight mercaptans before the charge ishydrogenated. In addition, and more for the purpose of eliminatingcorrosion in product stabilizer 45 and splitter 48, hydrogen sulfide maybe extracted from the unstabilized dripolene passing through line 29 bysuitable basic materials, of which mention may be made of sodium orpotassium hydroxide, monoethanolamine-water, diethanolamine-water, orsolid sodium carbonate.

Numerous other modifications may be made to the two principalembodiments described above, with the obtention of improved orequivalent results. For example, a portion of the hydrogen-containingrecycle gas may be heated out of the presence of the dripolene, and maybe added to the eifluent leaving Reaction Zone I and before passage intoReaction Zone H. Thus the low-catalyst requirements of the two-zoneembodiment is achieved while any possibility of coking up heater 18because of residual diolefins is of course obviated entirely.

#Eroirrthe discussion and'the examples-above, it isevidentlthatourcproce'ss is: :an extremely valuable improvement in thecatalytic: hydrogenation of olefinic hydrocarbonsintdriipolenefractions. By employing: our invention and contacting the dripolenewith-4t; platinum-alumina catalyst in the, presencezofi low: H 8 contenthydrogen gas while the'ditipoleneis .initiallyxsubstantiallyqentirely inthe liquid phase, it. is poss'ible to minimize or: reduce almostentirely the. quantity of: coke: formation heretofore exp erienced underthe prior..art: processes. .Our discovery mayfiheusedeithenin the: formof a single adiabatic reactiorr zone-.err-iployingsalowz space velocityor. may have twotorimone adiabatic zones with interstage heatingwherebysuhstantiah economies are achievecl'in respecttothenecessaryaquantity;of:catalyst.. .Having described the; invention, weclaim:

1. In a process for the catalytic selectivehydrogenatio'n of olefinic;hydrocarbons Lina a: low-boiling fraction, of dripolene, said dripolenebeing the normally liquid mixture of hydrocarbons obtained in thepyrolysis of normally gaseous hydrocarbons having at least two carbonatoms in the molecule at a temperature between about 1200 and 1800 'F.and a contact time between about 0.05 and 5 seconds, the improved methodof operation whereby conversion of olefins to coke is substantiallyreduced which comprises commingling said dripolene fraction with ahydrogen containing gas, said hydrogen containing gas having less than12 grains of hydrogen sulfide per standard cubic feet, and passing saidcommingled dripolene fraction and hydrogen-containing gas stream, whilesaid dripolene fraction is initially substantially in the liquidphaseandat a temperature between 100 and F into at least one fixedsubstantially-adiabatic bed of platinum-alumina hydrogenation catalyst,whereby monoolefins and diolefins in said dripolene fraction arehydrogenated and wherein said stream temperature is increased by theexothermic heat of hydrogenation to vaporize said dripolene fraction.

2. Process of claim 1 wherein said hydrogen-containing gas stream isemployed in a proportion of between about 500 and 10,000 standard cubicfeet per barrel of dripolene fraction, and the total bed inlet pressureis within the range of 100 to 1000 p.s.i.g.

3. Process of claim 1 wherein the hydrogen-containing gas stream isemployed within a proportion of between about 1000 and 4000 standardcubic feet per barrel of dripolene fraction, and the total bed inletpressure is within the range of 300 to 500 psig.

4. Process of claim 1 wherein the platinum-alumina hydrogenationcatalyst is spent hydroforming catalyst.

5. In a process for the catalytic selective hydrogenation of olefinichydrocarbons in a low-boiling fraction of dripolene, said dripolenebeing the normally liquid mixture of hydrocarbons obtained in thepyrolysis of normally gaseous hydrocarbons having at least two carbonatoms in the molecule at a temperature between about 1200- 1800 F. and acontact time between about 0.05 and 5 seconds, the improved method ofoperation whereby c0nversion of olefins to coke is substantially reducedwhich comprises commingling said dripolene fraction with ahydrogen-containing gas, said gas having less than 12 grains of hydrogensulfide per 100 standard cubic feet, passing said commingled dripolenefraction and hydrogen-containing gas while said dripolene fraction isinitially substantially in the liquid phase and at a temperature between100 and 150 F. into a first fixed substantiallyadiabatic bed ofplatinum-alumina hydrogenation catalyst, wherein monoolefins anddiolefins in said dripolene fraction are at least partially hydrogenatedand wherein said stream temperature is increased by the exothermic heatof hydrogenation to at least partially vaporize said dripolene fraction,withdrawing said commingled stream from the first bed and heating saidstream to a temperature within the range of about 300 F. to about 600F., and passing the heated commingled stream into at least 1 1 oneadditional fixed substantially-adiabatic bed of p1ati hum-aluminahydrogenation catalyst for additional hydrogenation of rnonoolefins anddiolefins.

6. Process of claim 5 wherein the hydrogen-containing gas stream isemployed in a proportion between about 500 and 10,000 standard cubicfeet per barrel of dripolene fraction, and the comrningled dripolenefraction and hydrogen-containing gas stream initially contacts the bedof hydrogenation catalyst at a total bed inlet pressure within the rangeof 100 to 1000 p.s.i.g.

7. Process of claim 5 wherein the hydrogen-containing gas stream isemployed in a proportion of between about 1000 and 4000 standard cubicfeet per barrel of dripolene fraction, and the commingled dripolenefraction and hydrogen-containing gas stream contacts the first bed ofhydrogenation catalyst at a total bed inlet pressure within the range of300 to 500 p.s.i.g.

8. Process .of claim 5 wherein the product stream after the last bed'ofhydrogenation catalyst is cooled to sepa rate a liquid product from anormalrgaseous material, and the liquid product. isthereafterfractionally distilled toseparate a relatively low-boilingrf'ractionforaromatics extraction from a relatively high boiling fraction for use asa motor fuel component.

9. Process'of'claim 5 wherein theplatinum-alumina hydrogenation catalystis spent hydroforming catalyst.

References Cited in the file of this patent UNITED STATES PATENTS UNITEDSTATES PATENT OFFICE CERTIFICATE OF CORRECTION Patent No. 2,953,612September 20 1960 Manford R, Haxton et al.,

It is hereby certified that error appears in the printed specificationof the above numbered patent requiring correction and that the saidLetters .Patent should read as corrected below.

Column 4, line 62 after "charge" insert sto k -=--g column 10, line 44,.for the claim reference numeral 1 7 read 2 Signed and 4th day of April1961.,

(SEAL) Attest: ERNEFH W. SWIDER WPQXMDW ARTHUR w. CROCKER AttestingOfl'icer Acting Commissioner of Patents

1. IN A PROCESS FOR THE CATALYTIC SELECTIVE HYDROGENATION OF OLEFINICHYDROCARBONS IN A LOW-BOILING FRACTION OF DRIPOLENE, SAID DRIPOLENEBEING THE NORMALLY LIQUID MIXTURE OF HYDROCARBONS OBTAINED IN THEPRYOLYSIS OF NORMALLY GASEOUS HYDROCARBONS HAVING AT LEAST TWO CARBONATOMS IN THE MOLECULE AT A TEMPERATURE BETWEEN ABOUT 1200 AND 1800*F.AND A CONTACT TIME BETWEEN ABOUT 0.05 AND 5 SECONDS THE IMPROVED METHODOF OPERATION WHEREBY CONVERSION OF OLEFINS TO COKE IS SUBSTANTIALLYREDUCED WHICH COMPRISES COMMINGLING SAID DRIPOLENE FRACTION WITH AHYDROGEN CONTAINING GAS, SAID HYDROGEN CONTAINING GAS HAVING LESS THAN12 GRAINS OF HYDROGEN SULFIDE PER 100 STANDARD CUBIC FEET, AND PASSINGSAID COMMINGLED DRIPOLENE FRACTION AND HYDROGEN-CONTAINING GAS STREAM,WHILE SAID DRIPOLENE FRACTION IS INITIALLY SUBSTANTIALLY IN THE LIQUIDPHASE AND AT A TEMPERATURE BETWEEN 100 AND 150* F., INTO AT LEAST ONEFIXED SUBSTANTIALLY-ADIABATIC BED OF PLATINUM-ALUMINUM HYDROGENATIONCATALYST, WHEREBY MONOOLEFINS AND DIOLEFINS IN SAID DRIPOLENE FRACTIONARE HYDROGENATED AND WHEREIN SAID STREAM TEMPERATURE IS INCREASED BY THEEXOTHERMIC HEAT OF HYDROGENATION TO VAPORIZE SAID DRIPOLENE FRACTION.